Naphtha Production via Catalytic Hydrotreatment of Refined Residual Lipids: Validation in Industrially Relevant Scale

Abstract

At the moment, there are no available data or studies exploring the production of naphtha boiling range hydrocarbons via hydroprocessing of pretreated residual lipids. To that aim, this study targets the production of naphtha, jet and diesel boiling range hydrocarbons via hydroprocessing of refined waste cooking oils utilizing solar hydrogen. The technology was first optimized in a TRL-3 plant. A heteroatom removal catalyst and a saturation catalyst were combined with an isomerization and hydrocracking catalyst to upgrade lipids. The results show that the severity of the process plays an important role in the yields of the fuels. Higher naphtha yields were observed at 663 K, 13.78 MPa and a liquid hourly space velocity of 0.33 h⁻¹, leading to the production of a fuel consisting of 34 wt% naphtha, 23 wt% jet and 42 wt% diesel boiling range hydrocarbons. Subsequently, the technology was validated and demonstrated in an industrially relevant unit (TRL-5). The results from the fuel characterization show that the diesel fraction can be used as a high-quality road transport drop-in fuel, as it is characterized by a high cetane index (~96) and a high flash point (414 K). Although jet and naphtha meet most commercial fuel specifications, further optimization of the process is necessary to meet fuel standards. In conclusion, the current work provides novel data relevant to industrial applications for road, aviation and maritime fuel production via hydroprocessing of refined waste cooking oil.

1. Introduction

Fuels are interlinked with growth and development in various sectors, like industry, agriculture, technology, etc. However, the increasing demand for fossil fuels is also linked with negative implications, such as significant fluctuation of fuel prices, dependence on fuel-producing countries, reduction of natural resources and the greenhouse effect and climate change. Green fuels could be part of the solution, as they can boost the national economy by creating new employment opportunities while contributing to the reduction of atmospheric pollution [1].

Hydroprocessing of lipid-based feedstocks for advanced diesel production and sustainable aviation fuels (SAFs) is a mature technology with commercial applications. The advanced diesel- and jet-like fuels from hydroprocessing of lipids are high-quality paraffinic fuels with favorable characteristics. Various types of lipid-based feedstocks have been explored, such as fresh vegetable oils [2,3], residual lipids (Waste Cooking Oils (WCOs)) [4], microbial oils [5,6], animal fats [7], microalgae oils [8,9] etc. In particular, the production of paraffinic fuels from fresh vegetable oils, animal fats and WCOs have reached commercial application. However, WCOs are highly heterogeneous, containing a variety of impurities that prevent their commercial exploitation, rendering a pretreatment stage essential. The pretreatment process depends on the lipid’s characteristics, the target products and the valorization routes.

There are many companies today that apply the aforementioned technology for advanced fuel production, such as Catalytic Condensation (UOP), NexOctane (Neste/KBR), Axens etc. [1,10]. The produced advanced diesel is a paraffinic fuel with a high cetane number (>77), high heating value (>44 MJ/kg) and low sulfur content (<5 wppm), which follow the EN 15940 specifications for paraffinic diesel-like fuels [11]. Studies on the effect of advanced paraffinic diesel-like fuels on internal combustion diesel engines reported that they can lead to the reduction of particulate matter, NOx, CO, CO₂ and total hydrocarbon emissions [10,11,12]. In addition, the advanced jet fuel produced via hydroprocessing of lipid-based feedstocks is also characterized by very good properties that comply with Jet A1 specifications. There are various studies that have tested these types of fuels on turbine engines [13].

Although hydroprocessing of lipid-based feedstocks for advanced paraffinic diesel and jet fuels is a mature technology available at commercial scale [10,14], the production of naphtha boiling range hydrocarbons has not been rigorously investigated [15]. Naphtha boiling range hydrocarbons (340–493 K) are lighter compared to those of jet (493–573 K) and diesel (573–633 K). Therefore, hydroprocessing of lipid-based feedstocks targeting naphtha boiling range hydrocarbons requires more severe operating conditions (higher temperatures and pressures) combined with isomerization, dewaxing and hydrocracking catalysts. At the moment, there are no available data or studies exploring the production of naphtha boiling range hydrocarbons via hydroprocessing of pretreated residual lipids.

To that aim, the current work fills the literature gap in this research field and presents the production of naphtha boiling range hydrocarbons via hydroprocessing of refined WCOs. The most common method for estimating the maturity of a technology is the Technology Readiness Level (TRL). This study is divided into three main parts. The first part includes the process optimization for the production of naphtha boiling range hydrocarbons via hydrotreatment of refined WCOs in a small-scale hydrotreatment (HDT) plant (TRL-3) and provides useful operational data on the current field. The second part involves technology validation and demonstration in an industrial-scale hydrotreatment plant (TRL-5), leading to the production of high amounts of organic liquid products. The final part includes the fractionation of the total organic liquid product, separating the naphtha, jet and diesel fractions for quality evaluation. Furthermore, the hydrogen for the hydrotreatment of WCOs was produced via water electrolysis utilizing solar energy, and therefore, the produced fuels can be called electrified fuels or e-fuels [16]. The current research provides novel and important data regarding industrial-scale application for electrified naphtha production via hydrotreatment of refined WCOs.

2. Methodology

2.1. Feedstock-WCO

WCOs, consisting mainly of triglycerides and fatty acids, were collected from restaurants and households [17]. A three-stage experimental process was developed to refine WCOs by removing solids, free fatty acids and water content. WCOs of different origins (households, restaurants, etc.) and quality were utilized, targeting their qualitative improvement. At the first stage (neutralization), the WCOs reacted with NaOH, forming water and soaps. The second stage involved the washing and removal of the soaps, leading to the third stage—the drying process. The acidity (Total Acid Number, TAN) and the quantity of the soaps were analyzed according to the ΕΝ 14104 [18] and AOCS Cc 17-95 methods [19], respectively, while the water content was determined with the ASTM D 6304 [20]. The process is described in more detail in the authors’ previous work [21]. The properties of the WCOs before and after pretreatment are presented in Table 1. The quality of the WCOs was notably enhanced, as indicated by the decrease in acidity and water content.

Table 1. Properties of WCOs before and after pretreatment.

Properties	Units	WCOs	Refined WCOs
Density at 288 K	g/ml	0.921	0.922
Sulfur	wppm	23.9	14.4 (523 1)
Hydrogen	wt%	11.92	11.87
Carbon	wt%	77.64	77.47
Nitrogen	wppm	165.5	26.3 (46 1)
Oxygen	wt%	10.4	10.6
Water dissolved	wt%	0.22	0.06
TAN	mgKOH/g	7.53	0.45
Viscosity at 313 K	cSts	34.42	34.13
H/C ratio	-	0.153	0.152
O/C ratio	-	0.133	0.133
HHV	MJ/kg	41.06	41.09
Pour point	K	-	267
CFPP	K	299	293
Simulated Distil. Curve
IBO	K	818	693
10 wt%	K	819	847
30 wt%	K	877	878
50 wt%	K	884	885
70 wt%	K	888	888
90 wt%	K	890	891
95 wt%	K	892	897
FBP	K	992	998

¹ Τhe values in brackets refer to the values after the addition of DMDS and TBA on refined WCO.

Table 1. Properties of WCOs before and after pretreatment.

Properties	Units	WCOs	Refined WCOs
Density at 288 K	g/ml	0.921	0.922
Sulfur	wppm	23.9	14.4 (523 1)
Hydrogen	wt%	11.92	11.87
Carbon	wt%	77.64	77.47
Nitrogen	wppm	165.5	26.3 (46 1)
Oxygen	wt%	10.4	10.6
Water dissolved	wt%	0.22	0.06
TAN	mgKOH/g	7.53	0.45
Viscosity at 313 K	cSts	34.42	34.13
H/C ratio	-	0.153	0.152
O/C ratio	-	0.133	0.133
HHV	MJ/kg	41.06	41.09
Pour point	K	-	267
CFPP	K	299	293
Simulated Distil. Curve
IBO	K	818	693
10 wt%	K	819	847
30 wt%	K	877	878
50 wt%	K	884	885
70 wt%	K	888	888
90 wt%	K	890	891
95 wt%	K	892	897
FBP	K	992	998

¹ Τhe values in brackets refer to the values after the addition of DMDS and TBA on refined WCO.

2.2. Catalyst

A commercial catalytic system was explored utilizing four different catalysts. The loading plan of the reactor starting from the top to bottom is as follows: 10 v/v % HDO (hydrodeoxygenation) catalyst, 30 v/v % HDT-saturation (hydrotreating) catalyst, 10 v/v % HDO catalyst (similar to the first zone), 40 v/v % dewaxing–isomerization catalyst and 10 v/v % hydrocracking catalyst. As the catalysts were commercial, no further details for the consistency and structure can be provided. In order to maintain a desired Liquid Hourly Space Velocity (LHSV), the catalysts were diluted with glass beads. Catalyst presulfiding was performed by a procedure defined by the catalyst manufacturer, utilizing LAGO (Light Atmospheric Gas Oil) with DMDS. DMDS (dimethyl disulfide) and tetra-butyl-amine (TBA) were added to the WCOs in order to increase the sulfur and nitrogen content, respectively, according to the catalyst provider, to keep the activity of the catalyst steady.

2.3. Analysis

The liquid products were collected daily and analyzed at CERTH’s analytical laboratory. The gaseous products were analyzed online via an Agilent 7890 GC and a 5975C mass spectrometer (Agilent, St. Clara, CA, USA). The methods are presented in Table 2.

Table 2. Analysis and methods for liquid samples.

Analysis	Methods	Accuracy	References
Density	ASTM D-4052	0.0002 g/ml	[22]
SIM-DIS	ASTM D-7169	4 °C	[23]
C content	LECO ASTM D-5291	0.60 wt%	[24]
H content	LECO ASTM D-5291	0.20 wt%	[24]
S content	XRFS analyzer ASTM D-4294	3 wt%	[25]
N content	ASTM D-4629	3 wt%	[26]
Water content	ASTM D-6304 or ASTM E-203	5 and 3 wt%	[20,27]
TAN	ASTM D-664	5%	[28]
Kinematic viscosity	ASTM D445	0.3%	[29]
Cetane index	ASTM D-976	Calculated by SIMDIS	[30]
Pour point	ASTM D-97	3 °C	[31]
Micro carbon residue (MCR)	ASTM D4530	1%	[32]
CFPP	EN116	1 °C	[33]

Table 2. Analysis and methods for liquid samples.

Analysis	Methods	Accuracy	References
Density	ASTM D-4052	0.0002 g/ml	[22]
SIM-DIS	ASTM D-7169	4 °C	[23]
C content	LECO ASTM D-5291	0.60 wt%	[24]
H content	LECO ASTM D-5291	0.20 wt%	[24]
S content	XRFS analyzer ASTM D-4294	3 wt%	[25]
N content	ASTM D-4629	3 wt%	[26]
Water content	ASTM D-6304 or ASTM E-203	5 and 3 wt%	[20,27]
TAN	ASTM D-664	5%	[28]
Kinematic viscosity	ASTM D445	0.3%	[29]
Cetane index	ASTM D-976	Calculated by SIMDIS	[30]
Pour point	ASTM D-97	3 °C	[31]
Micro carbon residue (MCR)	ASTM D4530	1%	[32]
CFPP	EN116	1 °C	[33]

The oxygen concentration was indirectly calculated through stoichiometric analysis (C, H) and measurements of sulfur and nitrogen content, assuming negligible concentrations of other elements in the liquid samples. A high heating value (HHV) was estimated according to the following equation, which was provided by [34] (Equation (1)):

HHV(MJ/kg) = 0.3491C + 1.1783H − 0.1034O − 0.0151N + 0.1005S − 0.0211ash

(1)

where C, H, O, N and S are the corresponding elemental compositions on a dry basis wt%. It was assumed that the ash content was zero in all samples.

2.4. Testing Infrastructure

Hydrotreatment optimization for the refined WCOs was carried out in a small-scale hydroprocessing plant (TRL-3). The unit consisted of a stainless steel continuous flow tubular reactor (with an inlet diameter of 15.8 mm and total length of 704 mm) containing six independent heating zones with a capacity ranging from 20 to 80 mL/h. This unit has been deeply described in the authors’ previous work [35]; therefore, no further details will be provided here.

After the optimization step, the technology was validated and demonstrated in an industrially relevant hydrotreatment plant at CERTH (TRL-5), which includes two stainless steel continuous flow tubular reactors in series (with an available volume of 555 cm³ each, 28 mm inside diameter and 955 mm length). Each reactor contains five independent heating zones. The current unit can accurately simulate operations from industrial units and can provide useful operational data for scaling up the process. More details can be found in the authors’ previous work [36]. As mentioned above, the hydrogen for both hydrotreatment pilot plants was produced via water electrolysis utilizing solar energy, leading to the production of electrified fuels or e-fuels [1,16].

Fractionation of the total organic liquid product obtained from the large-scale hydrotreatment plant (TRL-5) was performed on CERTH’s batch fractionation unit. This unit is a batch distillation unit with a minimum capacity of 10 L and a maximum capacity of 20 L per batch. It operates under vacuum conditions and can reach temperatures up to 637 K. The unit consists of a feed tank, with a capacity of 30 L, a fractionation column and a liquid product flask, with a capacity of 20 L.

2.5. Experimental Procedure

The first step of the current research involved the exploration and optimization of the examined technology on a small-scale hydrotreatment plant (TRL-3), targeting the maximization of naphtha yields. Five operating conditions were tested to examine the effects of reaction temperature, pressure and liquid hourly space velocity (LHSV) on naphtha yields and quality, as well as on process efficiency in terms of hydrogen consumption. The examined conditions are presented in Table 3. The testing operating window was selected after a detailed literature review screening triglyceride hydrotreatment [12,37,38,39,40,41,42]. According to the literature, a typical temperature range for triglyceride hydrotreatment is between 603 K and 663 K and a typical pressure range is between 10.34 MPa and 13.78 MPa, while in most studies the LHSV is close to 1 h⁻¹. To that aim, in condition 1, the lowest operating temperature and pressure were selected (603 K and 10.34 MPa) at LHSV 1 h⁻¹. In order to evaluate the effect of reaction temperature, in condition 2, the temperature was increased to 633 K, while all the other operating parameters were kept constant. In condition 3, the effect of LHSV was examined by reducing the LHSV from 1 h⁻¹ (condition 2) to 0.5 h⁻¹ (condition 3). The aim of condition 4 was to investigate the effect of operating pressure; therefore, the pressure increased from 10.34 MPa (condition 3) to 13.78 MPa (condition 4). Finally, we decided to explore if further reduction of the LHSV would improve hydrogen consumption. In condition 5, the LHSV was reduced to 0.33 h⁻¹ while all the other operating parameters were kept similar to those in condition 4. The stabilization of the catalyst performance under each condition was assessed by monitoring the sulfur content in the daily liquid products. When the sulfur content on the liquid product remained consistent over 2–3 consecutive days, the catalyst performance was considered stable, and the product sample was deemed representative of the current condition for comprehensive analysis. Additionally, the gaseous product was analyzed to facilitate mass balance and hydrogen consumption calculations.

Table 3. Testing operating window at small-scale hydrotreatment plant (TRL-3).

Parameters	Units	Cond. 1	Cond. 2	Cond. 3	Cond. 4	Cond. 5
Temperature	K	603	633	633	663	663
Pressure	MPa	10.34	10.34	10.34	13.78	13.78
Hydrogen/oil ratio	scfb	5000	5000	5000	5000	5000
LHSV	h−1	1	1	0.5	0.5	0.33
Duration	DOS 1	3	2	3	3	10

¹ DOS: days on stream.

Table 3. Testing operating window at small-scale hydrotreatment plant (TRL-3).

Parameters	Units	Cond. 1	Cond. 2	Cond. 3	Cond. 4	Cond. 5
Temperature	K	603	633	633	663	663
Pressure	MPa	10.34	10.34	10.34	13.78	13.78
Hydrogen/oil ratio	scfb	5000	5000	5000	5000	5000
LHSV	h−1	1	1	0.5	0.5	0.33
Duration	DOS 1	3	2	3	3	10

¹ DOS: days on stream.

The next step of the study involved the validation and demonstration of the process at a larger-scale hydrotreatment plant (TRL-5) to produce higher quantities of advanced e-fuels, using the selected operating conditions with the highest naphtha yields and the lowest hydrogen consumption. The produced advanced total liquid product from the large-scale hydrotreatment plant was inserted into the batch fractionation unit to separate the fractions of naphtha, jet and diesel. The three fractions were further analyzed and compared with the fossil fuel specifications of road transport diesel, gasoline and Jet A1.

3. Results

3.1. WCO Evaluation

WCOs originate from restaurants and households and mainly consist of triglycerides and fatty acids [17]. The fatty acid composition of the pretreated WCOs used in this study is presented in Figure 1. As can be observed, the WCOs consisted mainly of linoleic acid (C18:2) at 45.49 wt%, oleic acid (C18:1) at 36.63 wt%, palmitic acid (C16:0) at 10.45 wt% and stearic acid (C18:0) at 4.39 wt%. Based on this composition, the WCOs contained double bonds that had to be hydrocracked, leading to straight-chain paraffins. Furthermore, jet fuels must have a specific amount of aromatics because if the aromatic content of jet fuel is very low, the fuel will cause O-ring seals to shrink, harden and fail [43]. The high concentration of linoleic acid is significant, leading to final fuels with a high number of aromatics [44]. However, the high composition of aromatic compounds can lead to the erosion of turbine blades [45]. According to ASTM D7566 [46] for jet fuel, the aromatic content should range from 8 to 25 v/v %. Finally, lipids contain a high amount of oleic and stearic acids, which are important as they reduce mechanical losses during fuel flow through fuel systems.

Figure 1. Fatty acid composition of refined WCOs.

Figure 1. Fatty acid composition of refined WCOs.

According to the properties, WCOs have a higher density compared to the standards for commercial fossil fuels, including naphtha (0.720–0.775 gr/mL), jet (Jet A: 0.775–0.840 gr/mL) and diesel (road transport diesel: 0.820–0.845 gr/mL). The higher density is a result of the heavy hydrocarbon content of WCOs, with a boiling point starting from 653 K and ending at 1000 K. Therefore, a cracking step to break the heavy molecules and the double bonds to lighter straight-chain hydrocarbons is very important to produce advanced biofuels that can meet the specifications of commercial fuels. In addition, WCOs are characterized by high oxygen content that has to be removed via hydrodeoxygenation (HDO), decarbonylation and decarboxylation reactions. Furthermore, the viscosity of WCOs is higher compared to the commercial fuel standards of Jet A1 (<8.0 cSt) and road transport diesel EN 590 (2–4.5 cSt) [47]. Finally, the heating value has to be improved in order to meet the commercial diesel fuel specifications (EN590 for road transport diesel is 43.8 MJ/kg). The above analysis clearly shows that refined WCOs require an upgrading step, and catalytic hydrotreatment is a promising technology for this purpose.

3.2. Experimental Results from Process Optimization

The priority of WCO hydrotreatment is to remove the oxygen content. Three basic reactions are involved in oxygen removal, taking place at the same time; however, the rate of each reaction depends on the process operating conditions [11]. The most important reaction is HDO (Equation (2)), which is responsible for the two-phase liquid product that is produced after hydrotreatment. During HDO, the oxygen reacts with the hydrogen, leading to the formation of an aqueous phase, which is visually distinguished from the organic phase. At the beginning of these reactions, the carboxyl group reduces to a carbonyl group. During this reaction, fatty aldehydes are produced, which are then reduced further to form alcohols. These alcohols react with the hydrogen, leading to the formation of alkane or alkene and water. The alkane formation step is also called the hydrogenation mechanism, while the formation of alkene is called the dehydration step. Dissolved water content analysis, which was performed on the aqueous phase, showed that the water content was 99 wt%, indicating that the aqueous phase was mainly water.

Decarbonylation is the second type of reaction (Equation (3)) responsible for oxygen removal. During decarbonylation, the oxygen is removed in the form of water but also in the form of carbon monoxide (CO). However, carbon monoxide can react further with the hydrogen, leading to the formation of methane (CH₄) and water, according to the methanation reaction (Equation (4)), or methane and carbon dioxide (CO₂), according to equation 5. In the first step of the decarbonylation reaction, the carboxyl group is transformed into the carbonyl group. In the second step, oxygen is removed from the aldehyde in the form of carbon monoxide.

The third type of oxygen removal reaction is decarboxylation (Equation (6)), which lead to the formation of carbon dioxide (CO₂). However, the carbon dioxide can react further with the remaining hydrogen, forming methane and water via methanation reactions (Equation (7)) or leading to the formation of carbon monoxide and water via water gas shift reactions (Equation (8)).

R-CH₂-COOH + H₂ → R-CH₃ + CO + H₂O (decarbonylation)

(2)

R-CH₂-COOH → R-CH₃ + CO₂ (decarboxylation)

(3)

R-CH₂-COOH + 3H₂ → R-CH₂-CH₃ + 2H₂O (hydrodeoxygenation)

(4)

CO + 3H₂ <-> CH₄ + H₂O (methanation 1/3)

(5)

2CO + 2H₂ <-> CH₄ + CO₂ (methanation 2/3)

(6)

CO₂ + 4H₂ <-> CH₄ + 2H₂O (methanation 3/3)

(7)

CO₂ + H₂ <-> CO + H₂O (water gas shift)

(8)

Table 4 presents the gas product analysis results for the tested operating conditions (for condition No. 2, gas product analysis was not performed; the reason will be described below) and it can be observed that more than 85 wt% of the gas product was hydrogen (which did not react during the process). At a typical refinery hydrotreatment plant, the hydrogen contained in the gas product is recycled back to the reactor, and only the hydrogen that has reacted or been consumed is replaced with fresh hydrogen. However, the two pilot plants that were employed are once-through systems and do not have a recycling line; therefore, the hydrogen exits from the unit with the other gas products. According to the GC results, an increase in reactor temperature from 603 K (condition 1) to 633 K (condition 3) and a reduction in the retention time via the reduction of LHSV from 1 h⁻¹ (condition 1) to 0.5 h⁻¹ (condition 2) led to a gas product with no methane, but higher contents of propane, isobutane, n-butane, isopentane and n-pentane, indicating that more hydrocracking reactions took place, leading to higher amounts of gas hydrocarbons [37]. In general, the presence of gas hydrocarbons is not preferable, suggesting that there are carbon losses from the organic liquid product to the gas products. A further increase in the reactor temperature from 633 K (condition 3) to 663 K (condition 4), combined with an increase in the reactor pressure from 10.34 MPa (condition 3) to 13.78 MPa (condition 4), enhanced the hydrocracking reactions due to the more severe operating process, a phenomenon that is also observed during the hydrotreatment of petroleum-based feedstocks [48]. This change led to a three-fold increase in methane, indicating that methanation reactions are promoted by the severity of the process. In addition, from the comparison between condition 1 and condition 4, it can be observed that the carbon monoxide and carbon dioxide contents were reduced in condition 4 compared to 1. These decreases confirm the presence of more methanation reactions where the carbon monoxide and carbon dioxide reacted with hydrogen, leading to the formation of methane. Therefore, it can be concluded that increasing the severity of the process promotes methanation reactions, which are highly exothermic reactions and not preferable in industry. In addition, methanation reactions are also characterized by carbon losses in gas products, further verifying that these reactions should be avoided. Regarding condition 5, the liquid hourly space velocity was reduced further to 0.33 h⁻¹, leading to even higher retention time, and therefore, a more severe process. As expected, the carbon monoxide and carbon dioxide contents were reduced, and the methane content increased further. This result further validates that an increase in the process severity promotes methanation reactions but reduces carbon monoxide production, which is very toxic. In addition, the propane and n-butane contents increased, resulting in even higher carbon losses to the gas products. It was observed that the main components of the gas product were unreacted hydrogen, gas hydrocarbons (methane, propane ethane, butane, etc.) and small amounts of carbon monoxide and carbon dioxide. Similar studies on the hydroprocessing of lipid-based feedstocks, such as palm oil [42], soybean [49], animal fats [7] and waste cooking oil [40], have observed similar gas composition, confirming the results of the current research. In general, the aim is to reduce the production of gas hydrocarbons, as they represent carbon losses from the liquid product to the gas product. In addition, the minimization of carbon monoxide is very important, as it is highly toxic to human health.

Table 4. Gas product analysis of the five conditions at a small-scale hydrotreatment plant (TR-3).

Gases/wt%	Cond. 1	Cond. 2	Cond. 3	Cond. 4	Cond. 5
Hydrogen	90.705	-	90.423	86.641	85.371
Methane	1.045	-	0.000	3.998	4.981
Ethane	0.376	-	0.000	0.570	0.641
Propane	2.277	-	2.559	2.869	2.646
Isobutane	0.014	-	0.094	0.510	0.000
N-Butane	0.023	-	0.073	0.352	0.399
Isopentane	0.001	-	0.027	0.219	0.237
N-Pentane	0.003	-	0.018	0.119	0.116
Carbon Dioxide	1.304	-	0.000	0.767	0.619
Carbon Monoxide	0.055	-	0.000	0.067	0.061
Hydrogen Sulfide	0.029	-	0.000	0.024	0.020

Table 4. Gas product analysis of the five conditions at a small-scale hydrotreatment plant (TR-3).

Gases/wt%	Cond. 1	Cond. 2	Cond. 3	Cond. 4	Cond. 5
Hydrogen	90.705	-	90.423	86.641	85.371
Methane	1.045	-	0.000	3.998	4.981
Ethane	0.376	-	0.000	0.570	0.641
Propane	2.277	-	2.559	2.869	2.646
Isobutane	0.014	-	0.094	0.510	0.000
N-Butane	0.023	-	0.073	0.352	0.399
Isopentane	0.001	-	0.027	0.219	0.237
N-Pentane	0.003	-	0.018	0.119	0.116
Carbon Dioxide	1.304	-	0.000	0.767	0.619
Carbon Monoxide	0.055	-	0.000	0.067	0.061
Hydrogen Sulfide	0.029	-	0.000	0.024	0.020

Conversion and naphtha, jet and diesel selectivity for the products under the tested conditions, are presented in Figure 2 and were calculated using Equations (9)–(12):

Conversion (%) = (feed₆₃₃₊ − Product₆₃₃₊)/(Feed₆₃₃₊) × 100

(9)

Naphtha selectivity (%) = (Product_339–493 − Feed_339–493)/(Feed₄₉₃₊ − Product₄₉₃₊)

(10)

Jet selectivity (%) = (Product_493–573 − Feed_493–573)/(Feed₅₇₃₊ − Product₅₇₃₊)

(11)

Diesel selectivity (%) = (Product_573–633 − Feed_573–633)/(Feed₆₃₃₊ − Product₆₃₃₊)

(12)

The mild operating conditions promoted diesel production. In condition 1, the diesel, jet and naphtha selectivity were 87.5%, 11.5% and 1%, respectively, with a conversion rate of 96%. Increasing the operating temperature from 603 K (condition 1) to 633 K (condition 2) led to increases in jet (22.4%) and naphtha (3.6%) selectivity, while the diesel selectivity was reduced (74%); however, the conversion rate remained similar to that of condition 1 (at 96%). These results show that the increase in reaction temperature promoted cracking reactions, leading to the production of lighter boiling range hydrocarbons. An increase in retention time via the reduction of the LHSV from 1 h⁻¹ to 0.5 h⁻¹ led to higher jet (25.8%) and naphtha (4.1%) selectivity and lower diesel selectivity (70.1%) but also to a higher conversion rate (97%). Increasing the reaction time via the LHSV leads to the production of lighter boiling range hydrocarbons, which is expected as the reaction time for the feedstock is higher. Regarding more severe operation (condition 4), the pressure increased from 10.34 MPa (condition 3) to 13.78 MPa (condition 4) and the temperature from 633 K (condition 3) to 663 K (condition 4). It is noted that the naphtha selectivity (17.2%), as well as the jet selectivity (36.4%), increased since the more severe operation promoted hydrocracking reactions, leading to the production of lighter boiling range hydrocarbons. Finally, condition 5 aimed to maximize naphtha production; therefore, the LHSV was reduced further to 0.33 h⁻¹. Thus, the highest naphtha selectivity was achieved, reaching 34.3%; however, the selectivities for jet and diesel were reduced to 22.2% and 43.4%, respectively. These findings highlight that severe operation promotes hydrocracking reactions, leading to the production of lighter boiling range hydrocarbons. Therefore, the operating conditions must be wisely selected depending on the targeted products. The highest naphtha production was achieved in condition 5; however, the LHSV was reduced to a very low value, which means that in order to maintain the capacity of the unit, large reactors are required, increasing the investment cost of such plants. In addition, in all conditions, high conversion rates above 96% were achieved. Specifically, the highest conversion rate was observed in condition 5 (99%).

Figure 2. Naphtha, jet and diesel selectivity and conversion rate for the products of the tested conditions.

Figure 2. Naphtha, jet and diesel selectivity and conversion rate for the products of the tested conditions.

The mass recovery curve of both the refined WCOs and the products from the tested conditions are presented in Figure 3, showing that WCOs consist of heavy boiling range hydrocarbons, with 80 wt% of them having a boiling point above 873 K. In contrast, after the hydrotreatment process, the boiling ranges of the products were in the diesel, naphtha and jet boiling ranges. Particularly, the naphtha boiling range is from 339 K to 493 K, the jet range is from 493 K to 573 K and the diesel range is from 573 K to 633 K. The mass recovery plot confirms again that in conditions 1, 2 and 3, the naphtha boiling range hydrocarbons were very few, while the increase in process severity under conditions 4 and 5 resulted in higher naphtha yields. As the results between conditions 1 and 2 are similar, it was decided to not perform further analysis on the products of condition 2 (this is why there are no available data regarding the composition of the gas products under this condition in Table 4). The naphtha, jet and diesel yields under each condition are summarized below:

• Condition 1: Naphtha yields = 2 wt%, jet yields = 10 wt%, diesel yields = 84 wt%;
• Condition 2: Naphtha yields = 3 wt%, jet yields = 22 wt%, diesel yields = 71 wt%;
• Condition 3: Naphtha yields = 4 wt%, jet yields = 27 wt%, diesel yields = 66 wt%;
• Condition 4: Naphtha yields = 17 wt%, jet yields = 37 wt%, diesel yields = 45 wt%;
• Condition 5: Naphtha yields = 34 wt%, jet yields = 23 wt%, diesel yields = 42 wt%.

Figure 3. Mass recovery curves for the refined WCOs feedstocks and the hydrotreatment products under the tested conditions.

Figure 3. Mass recovery curves for the refined WCOs feedstocks and the hydrotreatment products under the tested conditions.

The properties of the products obtained under the tested conditions are presented in Table 5. Based on the results, it is evident that the density of the refined WCOs was reduced under all examined conditions below 0.80 g/mL, showing the effectiveness of hydrocracking reactions. In addition, the oxygen content was reduced to almost zero, reaching the detectable limits, which further confirms the effectiveness of the HDO, decarbonylation and decarboxylation reactions (as affirmed by the second aqueous phase in the products). Furthermore, the aqueous phase increased with the severity of the process operation, reaching the highest values in conditions 4 and 5. Furthermore, the sulfur content was reduced, showing the effectiveness of hydrodesulfurization (HDS) reactions. More specifically, regarding conditions 1, 2 and 4, the sulfur content in the liquid products was <10 wppm, which is the max limit according to commercial road transport diesel and gasoline fuel specifications. Another important observation is the reduction in viscosity, which was 42.6 cSt for the refined WCOs and reduced to less than 4 cSt for the examined products. The lowest viscosity value was observed in conditions 4 and 5, ascertaining that the severity of the process strongly influenced the viscosity of the final products. More intense operating conditions can lead to products with lower viscosity. The removal of heteroatoms also results in an increased heating value of the upgraded oils, especially as a consequence of their lower oxygen content [50]. The heating value of the refined WCOs was increased from 38.87 MJ/kg to more than 47 MJ/kg after hydrotreatment upgrading. This increase is in accordance with similar studies performed on waste cooking oil hydrotreatment [37]. Moreover, the cold flow properties were strongly influenced by the operating conditions; specifically, the severity of the process rendered products with a low cold filter plugging point (CFPP). Condition 5, where the retention time was the highest, presented the lowest CFPP value.

Hydrogen consumption of HVOs (hydrotreated vegetable oils), according to the literature [51], approaches 445 sl/L and is influenced by the level of saturation of the feedstock and the carbon chain length. In the current study, hydrogen consumption was calculated as standard liters of consumed hydrogen per liter of liquid feed (refined WCOs). The results from the tested conditions are in line with those in the literature and are very close to the hydrogen consumption of hydrotreated fresh vegetable oils (HVOs) [51]. According to the experimental results, it can be observed that the hydrogen consumption increased with the increase in process severity. This increase was anticipated as more hydrocracking reactions took place, producing lighter boiling range hydrocarbons. In general, it can be deduced that the organic liquid product via hydrotreatment of the refined WCOs presented good properties consisting of naphtha, jet and diesel boiling range hydrocarbons. Likewise, the operating conditions should be carefully selected, as they affect the yields of the produced hydrocarbons. If the target is the production of diesel range hydrocarbons, mild operation is preferable, while in the case of naphtha boiling range hydrocarbons, higher severity operation should be selected. In the next phase of this study, the examined technological approach was scaled up to a TRL-5 hydrotreatment plant to produce higher quantities of organic liquid product, which could then be fractionated into different fuel fractions (naphtha, jet and diesel) for quality characterization.

Table 5. Properties of hydrotreated WCOs under the tested conditions.

Properties	Units	Cond. 1	Cond. 2	Cond. 3	Cond. 4	Cond. 5
Density at 288 K	g/ml	0.789	0.787	0.786	0.777	0.747
Sulfur	wppm	4.94	5.12	22.50	2.70	12
Hydrogen	wt%	15.05	-	15.04	15.08	14.90
Carbon	wt%	84.95	-	84.9	84.89	85.05
Oxygen	wppm	0.00	-	0.06	0.03	0.05
Nitrogen	wt%	0.30	-	0.30	0.30	0.30
Water dissolved	wt%	0.001	-	0.001	0.001	0.004
Aqueous phase	v/v %	8.05	7.95	8.77	9.72	9.39
TAN	mgKOH/g	0.0	-	0.0	0.0	0.0
Viscosity at 313 K	cSts	3.853	-	3.454	2.425	1.352
HHV	MJ/kg	47.38	-	47.35	47.39	47.236
CFPP	K	294	-	292	267	255
Flash point	K	390	-	364	318	<284
Hydrogen consumption	Sl/l feed 1	351	-	401	417	420

¹ Standard liters of hydrogen per liter of liquid feed (refined WCO).

Table 5. Properties of hydrotreated WCOs under the tested conditions.

Properties	Units	Cond. 1	Cond. 2	Cond. 3	Cond. 4	Cond. 5
Density at 288 K	g/ml	0.789	0.787	0.786	0.777	0.747
Sulfur	wppm	4.94	5.12	22.50	2.70	12
Hydrogen	wt%	15.05	-	15.04	15.08	14.90
Carbon	wt%	84.95	-	84.9	84.89	85.05
Oxygen	wppm	0.00	-	0.06	0.03	0.05
Nitrogen	wt%	0.30	-	0.30	0.30	0.30
Water dissolved	wt%	0.001	-	0.001	0.001	0.004
Aqueous phase	v/v %	8.05	7.95	8.77	9.72	9.39
TAN	mgKOH/g	0.0	-	0.0	0.0	0.0
Viscosity at 313 K	cSts	3.853	-	3.454	2.425	1.352
HHV	MJ/kg	47.38	-	47.35	47.39	47.236
CFPP	K	294	-	292	267	255
Flash point	K	390	-	364	318	<284
Hydrogen consumption	Sl/l feed 1	351	-	401	417	420

¹ Standard liters of hydrogen per liter of liquid feed (refined WCO).

3.3. Demonstration and Validation of the Technology in Industrially Relevant Large-Scale Pilot Plant

The catalyst loading plan from the smaller (TRL-3) unit was adapted to match the larger-scale (TRL-5) plant; the main difference is that the larger unit consists of two reactors that operate in series. Therefore, it was decided to load the first reactor of the large unit with the HDO and saturation catalysts and the second reactor with the dewaxing and hydrocracking catalysts, according to the loading plan presented in Figure 4.

Figure 4. Catalyst loading plan for the reactors of the small-scale (TRL-3) and large-scale (TRL-5) units.

Figure 4. Catalyst loading plan for the reactors of the small-scale (TRL-3) and large-scale (TRL-5) units.

The objective was the maximization of naphtha boiling range hydrocarbons. Several operating conditions were examined to maximize the yields of naphtha, as presented in Table 6. The first condition was the base case condition, with mild activity, while the remaining conditions were determined during the experiment based on the simulated distillation results, with the goal of maximizing naphtha yields.

Table 6. Tested operating window at large-scale hydrotreatment plant (TRL-5).

Parameters	Temp. Reactor 1 (K)	Temp. Reactor 2 (K)	Pressure (MPa)	LHSV (h−1)	H2/Oil Ratio (Scfb)
Condition 1	603	603	10.34	0.40	5000
Condition 2	623	623	10.34	0.40	5000
Condition 3	633	633	10.34	0.35	5000
Condition 4	633	643	10.34	0.35	5000
Condition 5	573	573	10.34	0.6	5000
Condition 6	573	613	10.34	0.6	5000
Condition 7	573	633	10.34	0.5	5000
Condition 8	573	633	10.34	0.3	5000

Table 6. Tested operating window at large-scale hydrotreatment plant (TRL-5).

Parameters	Temp. Reactor 1 (K)	Temp. Reactor 2 (K)	Pressure (MPa)	LHSV (h−1)	H2/Oil Ratio (Scfb)
Condition 1	603	603	10.34	0.40	5000
Condition 2	623	623	10.34	0.40	5000
Condition 3	633	633	10.34	0.35	5000
Condition 4	633	643	10.34	0.35	5000
Condition 5	573	573	10.34	0.6	5000
Condition 6	573	613	10.34	0.6	5000
Condition 7	573	633	10.34	0.5	5000
Condition 8	573	633	10.34	0.3	5000

The resulting mass recovery curves of the organic liquid products are presented in Figure 5. As noted, only conditions 3 and 4 were able to provide a product with more than 15 wt% naphtha boiling range hydrocarbons. In condition 3, they made up 19 wt%, and in condition 4, they made up 17 wt%. Conditions 3 and 4 were the most severe among those examined, as the retention time was the highest, with an LHSV at 0.35 h⁻¹. However, it was found that the product losses under these conditions (3 and 4) were very high, at 9.6 wt% and 7.8 wt% respectively. In addition, due to the very low LHSV (0.35 h⁻¹), the feed flow was also very low, and the production of large product quantities was impossible in an acceptable experimental time.

Figure 5. Mass recovery curves for the products under the tested conditions from the large-scale (TRL-5) unit.

Figure 5. Mass recovery curves for the products under the tested conditions from the large-scale (TRL-5) unit.

These results highlight that the findings from the small-scale (TRL-3) hydrotreatment plant were not replicable at the larger-scale (TRL-5) plant. It was found that the main reason was that during hydrotreatment of the refined WCOs, the carbon monoxide, carbon dioxide and aqueous phase that were formed on the first reactor poisoned the active sites of the isomerization–dewaxing and hydrocracking catalysts of the second reactor, rendering limited naphtha and jet yields [52,53]. Unfortunately, in the TRL 5 plant, the two reactors operate in series at the same pressure, and there is no product separation system between them. As a result, all the products (gases and liquids) produced in the first reactor go to the second reactor. Therefore, the by-products (aqueous phase, CO and CO₂) from the reactions in the first reactor poison the catalyst active sites of the second reactor. To overcome this process limitation, it was decided to proceed and perform the hydrotreatment of refined WCOs in a two-step process. In the first step, the two reactors were filled only with HDO, HDS, HDN and saturation catalysts in order to remove the heteroatoms from the feed, leading to a two-phase liquid product consisting of an organic phase and an aqueous phase. After this step, the organic phase was separated from the aqueous phase via gravity and collected to be the main feedstock for the second hydrotreatment step. In the second step, the two reactors were loaded with dewaxing and hydrocracking catalysts. Therefore, neither the aqueous phase nor the gas products from the first step passed through the dewaxing and hydrocracking catalysts, allowing the catalysts to maintain their activity at high rates. The catalyst loading in the two-step process is presented in Figure 6.

Figure 6. Catalyst loading plan for the large-scale (TRL 5) plant in a two-step hydrotreatment.

Figure 6. Catalyst loading plan for the large-scale (TRL 5) plant in a two-step hydrotreatment.

The first step lasted 42 days on stream and ended due to limited availability of the refined WCOs. During this step, 405 L of organic liquid product was produced. During the 42 days of continuous operation, the catalyst activity was monitored by tracking increases in temperature, based on the sulfur content of the liquid products, to produce an upgraded product with low sulfur content that met the targeted specifications (S content <10 wppm). During the experiment, the temperature was increased manually over the days on stream to keep the sulfur content of the products stable at low levels (<150 wppm) (Figure 7). No DP or coke formation was observed in the unit.

Figure 7. Temperature variation over DOS on the left side (orange color line), and sulfur content variation over DOS on the right axis (blue color line) during the first step. Blue dots present the exactly value of each DOS while the black line shows the EN 590 specification for sulphur content (10 wppm) for the commercial market diesel.

Figure 7. Temperature variation over DOS on the left side (orange color line), and sulfur content variation over DOS on the right axis (blue color line) during the first step. Blue dots present the exactly value of each DOS while the black line shows the EN 590 specification for sulphur content (10 wppm) for the commercial market diesel.

The mass recovery curves (distillation curves) for the organic liquid products from some randomly chosen days are presented in Figure 8. It can be observed that, after the first hydrotreatment step with heteroatom removal and saturation catalysts, the heavy boiling range feedstock of refined WCOs was transformed to lower boiling point hydrocarbons in the diesel range. In addition, the oxygen content in the products was less than 2 wt%, which was due to the strong HDO, decarbonylation and decarboxylation reactions performed during the first step. The total organic liquid product from all days (DOS) of the first step was collected for the second step of isomerization–dewaxing and hydrocracking.

Figure 8. Mass recovery curves for randomly selected DOS products from the first-step process at the TRL5 hydrotreatment pilot plant.

Figure 8. Mass recovery curves for randomly selected DOS products from the first-step process at the TRL5 hydrotreatment pilot plant.

After the second hydrotreatment step, the resulting product contained naphtha, jet and diesel boiling range hydrocarbons, as presented in Figure 9 (random products). In order to maintain constant catalyst activity, the temperature of the reactors was increased manually over days on stream (DOS), as presented in Figure 10. The main goal was to keep the sulfur content of the products below 10 wppm, which is the maximum limit in commercial fossil diesel and gasoline specifications. In the next phase, the total organic liquid product from all days was collected and fractionated to separate the three different fractions (naphtha, jet and diesel) for quality evaluation.

Figure 9. Mass recovery curves for randomly selected DOS products from the second hydrotreatment step.

Figure 9. Mass recovery curves for randomly selected DOS products from the second hydrotreatment step.

Figure 10. Temperature variation over DOS on the left side (blue color line), and sulfur content variation over DOS on the right axis (red color line) during the second step.

Figure 10. Temperature variation over DOS on the left side (blue color line), and sulfur content variation over DOS on the right axis (red color line) during the second step.

The total mass balance from the TRL5 unit via the two-step process is presented in Figure 11. As expected, the mass balance closure was not 100% because the described unit was a TRL 5 pilot plant with high liquid and gas flows. However, the mass closure in stage 1 was 95.2%, while in stage 2, it was 96.9%, which is very high percentage for a pilot plant at this TRL level. According to the mass balance results, it can be observed that the mass yields from the fractionation of the final organic liquid product do not match the results from the TRL3 pilot plant. Regarding the TRL 3 results, the total liquid product consisted of 34 wt% naphtha, 23 wt% jet and 42 wt% diesel boiling range hydrocarbons; however, regarding the TRL5 results, the final liquid fuel consisted of 24.90 wt% naphtha, 17.86 wt% jet and 57.24 wt% diesel boiling range hydrocarbons. Although the yields for naphtha were reduced by scaling up the technology in TRL5, they remained high: from 1 kg of lipids, 176.86 g naphtha, 126.85 g jet and 406.56 g diesel boiling range hydrocarbons were produced. The results from the research show that the technology was successfully scaled up from the TRL3 to the TRL5 pilot plant; however, further optimization is required for the TRL 5 process.

Figure 11. Mass balance of the TRL 5 HDT plant, including both stages.

Figure 11. Mass balance of the TRL 5 HDT plant, including both stages.

3.4. Fuel Quality Evaluation

After the fractionation of the total organic liquid product, the three fractions of naphtha, jet and diesel were analyzed and compared with the fuel specifications (the boiling range of diesel is from 573 K to 633 K when the jet fraction is not included; when the jet fraction is included, the range is from 493 K to 633 K). According to the mass recovery curve presented in Figure 12, the diesel fraction consisted of 100 wt% hydrocarbons in the diesel boiling range, indicating that the fractionation of the diesel boiling range hydrocarbons was performed successfully. The composition of the diesel fraction was mainly n-paraffins (58.4 wt%) and iso-paraffins (41.3 wt%) (Figure 13).

Figure 12. Mass recovery curve of naphtha, jet and diesel fractions of advanced e-fuels.

Figure 12. Mass recovery curve of naphtha, jet and diesel fractions of advanced e-fuels.

Figure 13. Results from GC analysis performed on naphtha, jet and diesel fractions.

Figure 13. Results from GC analysis performed on naphtha, jet and diesel fractions.

As WCOs quantities are not adequate to completely replace fossil fuels via the production of e-fuels, they can be utilized as drop-in fuels. Therefore, blends of 10 v/v % e-fuel with conventional fuels were prepared and evaluated.

Regarding the diesel fraction, the properties of diesel specifications (limits), commercial diesel, e-diesel and a blend with 90 v/v % commercial diesel and 10 v/v % e-diesel are presented in Table 7. According to the analysis, the e-diesel fraction met most commercial diesel specifications, with the exceptions of density (0.7878 g/mL), which was lower compared to the limit (0.8200 g/mL–0.8450 g/mL); CFPP (286 K) which was higher compared to the limit (267 K is the limit for summer-grade diesel); and lubricity, which was 620 μm compared to the limit of 460 μm. The lower density is the result of the paraffinic nature of the e-diesel fraction (see Figure 13). CFPP is the lowest temperature at which a fuel can flow through a standard test filter under specified conditions, providing a more precise measurement that serves as a better indicator of fuel performance in an engine. The CFPP was higher for the e-diesel fraction compared to the limit due to the paraffinic nature of the fuel. However, this drawback can be overcome if a higher amount of isomerization catalyst is added during the process. Finally, the lubricity was higher compared to the specifications. In general, aromatics in fuels can contribute to lubricity, whereas paraffinic fuels, which have low aromatic content, generally exhibit poorer lubricity. Additionally, traditional fuels with higher sulfur levels tend to have better lubricity because sulfur compounds act as natural lubricants. Ultra-low sulfur diesel (ULSD), which has low sulfur levels, often requires additives to restore lubricity. The e-diesel fraction, which had almost zero sulfur content, had low lubricity. Therefore, lubricity-improving additives have to be added to enhance the lubricity of e-diesel to meet the fuel standards. Although only the drawbacks of e-diesel have been presented so far, the e-diesel fraction also presents some benefits compared to commercial diesel. First of all, it presents a much higher cetane index (96.7) compared to commercial diesel (56.9). The cetane index is a critical indicator of diesel fuel quality that directly affects the ignition delay and the combustion efficiency. A higher cetane index is preferred for smooth, efficient engine operation and low emissions. On the other side, a lower cetane index indicates a longer ignition delay, which can result in rough engine operation, noise and high emissions [54]. The flash point indicates how easily a fuel can ignite under normal handling and storage conditions. Diesel’s relatively high flash point makes it safer to store and handle compared to more volatile fuels, like gasoline. For safety reasons, the minimum flash point for diesel fuel is generally set around 325 K to ensure that it is not classified as a highly flammable liquid. Fuels with higher flash points are less hazardous and safer to transport. E-diesel presents a higher flash point compared to commercial diesel, which makes it safer than commercial diesel, as it is less prone to vaporize and ignite under normal conditions. As previously mentioned, a blend with 10 v/v % e-diesel and 90 v/v % fossil diesel was prepared and analyzed. The results are presented in Table 7. The addition of 10 v/v % e-diesel to fossil diesel improved some fuel characteristics of fossil diesel, such as the cetane index, the flash point, etc., while the blend met all fuel specifications for commercial diesel [10,55]. These results show that the e-diesel fraction could be potentially utilized as a drop-in fuel in commercial fossil diesel at rates up to 10 v/v %.

Table 7. Diesel specifications limits (for summer-grade), commercial diesel properties, e-diesel fraction properties and 90/10 v/v % commercial diesel/e-diesel fraction properties.

Properties	Units	Method	EN 590 Diesel	Commercial Diesel	e-Diesel	10% Blend
Density at 288 K	g/mL	EN ISO 12185 [56]	0.8200–0.8450	0.8267	0.7878	0.8228
Flash point	K	EN ISO 2719 [57]	>328	332	414	335
Total sulfur	ppm-w	EN ISO 20846 [58]	<10	7.8	0.4	7.2
Kinematic viscosity at 313 K	cSt	EN ISO 3104 [59]	2–4.5	2.728	3.981	2.927
Cetane index		EN ISO 4264 [60]	>46	56.9	96.7	60.1
Cetane number		EN ISO 5165 [61]	>51	57.2	>77.1	57.9
Carbon residue on 10% dist.res	% w/w	EN ISO 10370 [62]	<0.3	0.05	0.05	0.05
Copper strip corros.,3 h 223 K		EN ISO 2160 [63]	Class 1a	Class 1a	Class 1a	Class 1a
Water (K-F) in products	ppm-w	EN ISO 12937 [64]	<200	130	68	98
Ash	% w/w	EN ISO 6245 [65]	<0.01	0.005	0.004	0.005
C.F.P.P. 1	K	EN 116 [33]	>278	267	286	269
Total contamination	mg/Kg	EN 12662 [66]	<24	15.1	11.6	14.0
Oxidation stability	h	EN 15751 [67]	>20	>20	>20	>20
Polyaromatic hydrocarbons	% w/w	EN 12916 [68]	<8	3.0	0.6	1.7
Lubricity, c.w.s.d. 1.4 213 K	μm	EN ISO 12156-1 [69]	<460	320	620	340
95 v/v % recovered	K	EN ISO 3405 [70]	<633.1	631.2	603.2	630.5
Recovered at 523 K	v/v %	EN ISO 3405 [70]	<338.1	308.3	273.1	304.0
Recovered at 623 K	v/v %	EN ISO 3405 [70]	>358	365.8	>368.1	366.4

¹ C.F.P.P. spec is for summer-grade diesel.

Table 7. Diesel specifications limits (for summer-grade), commercial diesel properties, e-diesel fraction properties and 90/10 v/v % commercial diesel/e-diesel fraction properties.

Properties	Units	Method	EN 590 Diesel	Commercial Diesel	e-Diesel	10% Blend
Density at 288 K	g/mL	EN ISO 12185 [56]	0.8200–0.8450	0.8267	0.7878	0.8228
Flash point	K	EN ISO 2719 [57]	>328	332	414	335
Total sulfur	ppm-w	EN ISO 20846 [58]	<10	7.8	0.4	7.2
Kinematic viscosity at 313 K	cSt	EN ISO 3104 [59]	2–4.5	2.728	3.981	2.927
Cetane index		EN ISO 4264 [60]	>46	56.9	96.7	60.1
Cetane number		EN ISO 5165 [61]	>51	57.2	>77.1	57.9
Carbon residue on 10% dist.res	% w/w	EN ISO 10370 [62]	<0.3	0.05	0.05	0.05
Copper strip corros.,3 h 223 K		EN ISO 2160 [63]	Class 1a	Class 1a	Class 1a	Class 1a
Water (K-F) in products	ppm-w	EN ISO 12937 [64]	<200	130	68	98
Ash	% w/w	EN ISO 6245 [65]	<0.01	0.005	0.004	0.005
C.F.P.P. 1	K	EN 116 [33]	>278	267	286	269
Total contamination	mg/Kg	EN 12662 [66]	<24	15.1	11.6	14.0
Oxidation stability	h	EN 15751 [67]	>20	>20	>20	>20
Polyaromatic hydrocarbons	% w/w	EN 12916 [68]	<8	3.0	0.6	1.7
Lubricity, c.w.s.d. 1.4 213 K	μm	EN ISO 12156-1 [69]	<460	320	620	340
95 v/v % recovered	K	EN ISO 3405 [70]	<633.1	631.2	603.2	630.5
Recovered at 523 K	v/v %	EN ISO 3405 [70]	<338.1	308.3	273.1	304.0
Recovered at 623 K	v/v %	EN ISO 3405 [70]	>358	365.8	>368.1	366.4

¹ C.F.P.P. spec is for summer-grade diesel.

As depicted in Figure 12, the e-jet fraction consisted of 83 wt% jet range hydrocarbons, while naphtha and diesel boiling range hydrocarbons made up 12 wt% and 5 wt%, respectively. The properties of the e-jet fraction, the commercial jet and the blend with 10 v/v % e-jet fraction and 90 v/v % commercial jet are presented in Table 8 (for comparison purposes, the Jet A1 specifications are juxtaposed). The analysis results show that the e-jet fraction presented a lower density compared to the Jet A1 specifications due to its high paraffinic content. Furthermore, its distillation curve (10 v/v % recovered) exceeded the rate of Jet A1 specifications, but this could be addressed by changing the fractionation limits during the fractionation process. In addition, the total acidity was slightly higher compared to the Jet A1 specifications. Jet fuel acidity refers to the presence of acidic compounds in jet fuel, which can be harmful to fuel systems and components. It is typically measured to ensure the fuel meets quality and safety standards. The slightly higher TAN of the e-jet fraction can be explained by the low olefinic content of the fuel, which affects the acidity.

Another specification that was not met was the specific conductivity, which was 1 pS/m for the e-jet fraction compared to 50–600 pS/m in the Jet A1 specifications. Specific conductivity in jet fuel refers to the fuel’s ability to conduct electrical current, measured in picosiemens per meter (pS/m). It is a critical property because it affects the fuel’s susceptibility to static electricity buildup during handling and transfer processes. When jet fuel is pumped, filtered or transferred through pipelines, friction can generate static charges. If the specific conductivity is too low, the static charge cannot dissipate quickly, increasing the risk of electrostatic discharge, which can cause fires or explosions. Higher specific conductivity allows static charges to dissipate safely through the fuel to grounded surfaces. Jet fuel, in its natural state, has very low conductivity (around 1–5 pS/m) because it is a non-polar hydrocarbon liquid. Nevertheless, with the additions of static dissipator additives, the conductivity can be improved.

The key specification that remained unmet was the freezing point, which was 266 K for the e-jet fraction, while the specifications demand a freezing point below 226 K. The freezing point of jet fuel is the temperature at which wax crystals begin to form, causing the fuel to solidify or become partially gelled. This is a critical property because jet fuel must remain flowable and usable at the extremely low temperatures encountered at high altitudes. If the fuel begins to freeze, it can block fuel lines, filters and pumps, potentially causing engine failure. The freezing point of 226 K is used internationally and preferred for long-haul flights. Fuels with a higher concentration of paraffinic hydrocarbons, like the e-jet fraction, have higher freezing points because wax crystals form more easily. Aromatics and naphthenic hydrocarbons help to reduce the freezing point, but the e-jet fraction had no aromatic or naphthenic content. As the freezing point of the e-jet fraction did not meet the standard, it can be improved with additives or via blends with fuels like Jet B.

Overall, the e-jet fraction presented very good properties, such as a high flash point, low sulfur and aromatic contents and a high heating value, which are within the limits of the Jet A1 specifications. Specifically, the e-jet fraction contained n-paraffins (39.2 wt%) and iso-paraffins (59.9 wt%), as shown in Figure 13. The only drawbacks for the e-jet fraction are its freezing point, the specific conductivity and the small deviations in the distillation point. However, the blend of e-jet with fossil jet met almost all Jet A1 specifications, with the only exception of the freezing point, which was slightly higher (234 K) compared to the specifications (<226 K). This could possibly be overcome by blending the e-jet at a lower percentage with the commercial jet or with additives that can improve the freezing point. Finally, another way to improve the freezing point of the e-jet fraction is to apply an extra cyclization step in the production process to transform the paraffins into cyclo-paraffins, which are characterized by a lower freezing point.

Table 8. Properties of the e-jet fraction, the commercial jet fraction and their blend (Jet A1 specifications are also presented).

Properties	Units	Method	Jet A1 Specifications	Commercial Jet	e-Jet	Blend (90/10 v/v % Commercial Jet/e-Jet Fraction)
Density at 288 K	g/mL	ASTM D4052 [22]	0.7750–0.8400	0.7979	0.7743	0.7955
I.B.P.	K	ASTM D86 [71]		423.2	487.3	425.7
10 v/v % recovered	K	ASTM D86 [71]	<478.1	445.1	508.6	446.7
50 v/v % recovered	K	ASTM D86 [71]		472.2	531.9	475.8
90 v/v % recovered	K	ASTM D86 [71]		509.7	551.6	520.4
F.B.P.	K	ASTM D86 [71]	<573.1	527.4	560.5	539.9
Residue	v/v %	ASTM D86 [71]	<1.5	1.2	1.2	1.2
Loss	v/v %	ASTM D86 [71]	<1.5	0.4	0.4	0.4
Saybolt Color		ASTM D156 [72]		30	30	30
Appearance		VISUAL	Clear & Bright	Clear & Bright	Clear & Bright	Clear & Bright
Flash point (Tag)	K	ASTM D56 [73]	>311.1	314.1	367.1	315.1
Copper strip corros.,2 h-373 K		ASTM D130 [74]	Class 1a-Class 1b	Class 1a	Class 1a	Class 1a
Total sulfur	% w/w	ASTM D4294 [25]	<0.3	0.100	0.0003	0.096
Total acidity	mg KOH/g	ASTM D3243 [75]	<0.015	0.0026	0.0220	0.0030
Mercaptan Sulfur	% w/w	ASTM D3277 [76]	<0.003	0.0010	0.0003	0.0007
Existent gum	mg/100 mL	ASTM D381 [77]	<7		2.1	1.4
FIA/aromatics	v/v %	ASTM D1319 [78]	<25	15.9	1.5	15.6
Net heat of combustion	Mj/kg	ASTM D3338 [79]		43.31	44.06	43.35
Freezing point	K	ASTM D2386 [80]	<226	223	266	234
Kinematic viscosity at 253 K	cSt	ASTM D445 [29]	<8	4.200		1.297
Smoke point	mm	ASTM D1322 [81]	>18	23.0	>50	26.3
Test temperature	K	ASTM D3241 [82]	533.1	533.1	533.1	533.1
Change in pressure drop	mm Hg	ASTM D3241 [82]	<25	0	1	1
Interferometric Tube rating	nm	ASTM D3241 [82]	<85		18	21
Naphthalenes	v/v %	ASTM D1840 [83]	<3	0.80	0.06	0.69
Microseparometer rating (MSEP-A)		ASTM D3948 [84]	>70	87	92	90
Specific conductivity	pS/m	ASTM D2624 [85]	50–600	500	1	343
Particulate matter	mg/L	ASTM D5452 [86]	<1	0.34	0.05	0.30
Cum. channel particle counts ≥ 4 μm		IP 565/ISO 4406 [87,88]		149.0/14	176.1/15	131.6/14
Cum. channel particle counts ≥ 6 μm		IP 565/ISO 4406 [87,88]		48.0/13	84.9/14	10.3/11
Cum. channel particle counts ≥ 14 μm		IP 565/ISO 4406 [87,88]		2.0/08	24.6/12	0.9/07
Cum. channel particle counts ≥ 21 μm		IP 565/ISO 4406 [87,88]		1.0/06	11.3/11	0.3/06
Cum. channel particle counts ≥ 25 μm		IP 565/ISO 4406 [87,88]		0.0/06	6.3/10	0.2/05
Cum. channel particle counts ≥ 30 μm		IP 565/ISO 4406 [87,88]		0.0/04	3.2/09	0.1/05

Table 8. Properties of the e-jet fraction, the commercial jet fraction and their blend (Jet A1 specifications are also presented).

Properties	Units	Method	Jet A1 Specifications	Commercial Jet	e-Jet	Blend (90/10 v/v % Commercial Jet/e-Jet Fraction)
Density at 288 K	g/mL	ASTM D4052 [22]	0.7750–0.8400	0.7979	0.7743	0.7955
I.B.P.	K	ASTM D86 [71]		423.2	487.3	425.7
10 v/v % recovered	K	ASTM D86 [71]	<478.1	445.1	508.6	446.7
50 v/v % recovered	K	ASTM D86 [71]		472.2	531.9	475.8
90 v/v % recovered	K	ASTM D86 [71]		509.7	551.6	520.4
F.B.P.	K	ASTM D86 [71]	<573.1	527.4	560.5	539.9
Residue	v/v %	ASTM D86 [71]	<1.5	1.2	1.2	1.2
Loss	v/v %	ASTM D86 [71]	<1.5	0.4	0.4	0.4
Saybolt Color		ASTM D156 [72]		30	30	30
Appearance		VISUAL	Clear & Bright	Clear & Bright	Clear & Bright	Clear & Bright
Flash point (Tag)	K	ASTM D56 [73]	>311.1	314.1	367.1	315.1
Copper strip corros.,2 h-373 K		ASTM D130 [74]	Class 1a-Class 1b	Class 1a	Class 1a	Class 1a
Total sulfur	% w/w	ASTM D4294 [25]	<0.3	0.100	0.0003	0.096
Total acidity	mg KOH/g	ASTM D3243 [75]	<0.015	0.0026	0.0220	0.0030
Mercaptan Sulfur	% w/w	ASTM D3277 [76]	<0.003	0.0010	0.0003	0.0007
Existent gum	mg/100 mL	ASTM D381 [77]	<7		2.1	1.4
FIA/aromatics	v/v %	ASTM D1319 [78]	<25	15.9	1.5	15.6
Net heat of combustion	Mj/kg	ASTM D3338 [79]		43.31	44.06	43.35
Freezing point	K	ASTM D2386 [80]	<226	223	266	234
Kinematic viscosity at 253 K	cSt	ASTM D445 [29]	<8	4.200		1.297
Smoke point	mm	ASTM D1322 [81]	>18	23.0	>50	26.3
Test temperature	K	ASTM D3241 [82]	533.1	533.1	533.1	533.1
Change in pressure drop	mm Hg	ASTM D3241 [82]	<25	0	1	1
Interferometric Tube rating	nm	ASTM D3241 [82]	<85		18	21
Naphthalenes	v/v %	ASTM D1840 [83]	<3	0.80	0.06	0.69
Microseparometer rating (MSEP-A)		ASTM D3948 [84]	>70	87	92	90
Specific conductivity	pS/m	ASTM D2624 [85]	50–600	500	1	343
Particulate matter	mg/L	ASTM D5452 [86]	<1	0.34	0.05	0.30
Cum. channel particle counts ≥ 4 μm		IP 565/ISO 4406 [87,88]		149.0/14	176.1/15	131.6/14
Cum. channel particle counts ≥ 6 μm		IP 565/ISO 4406 [87,88]		48.0/13	84.9/14	10.3/11
Cum. channel particle counts ≥ 14 μm		IP 565/ISO 4406 [87,88]		2.0/08	24.6/12	0.9/07
Cum. channel particle counts ≥ 21 μm		IP 565/ISO 4406 [87,88]		1.0/06	11.3/11	0.3/06
Cum. channel particle counts ≥ 25 μm		IP 565/ISO 4406 [87,88]		0.0/06	6.3/10	0.2/05
Cum. channel particle counts ≥ 30 μm		IP 565/ISO 4406 [87,88]		0.0/04	3.2/09	0.1/05

Finally, the e-naphtha fraction consisted of n-paraffins (14 wt%), iso-paraffins (72.8), naphthene (7.2 wt%), olefins (1.7 wt%) and aromatics (1.6 wt%) (Figure 13). The properties of the commercial gasoline, e-naphtha, their blend (90 v/v % commercial gasoline with 10 v/v % e-naphtha) and the gasoline specifications (EN 228 [89]) are presented in Table 9. The fraction of e-aphtha did not meet some of the relative specifications, like the vapor pressure, octane number and distillation curve. Vapor pressure in gasoline refers to the pressure exerted by the gasoline’s vapors when it is in a closed container at a specific temperature. It is a measure of the fuel’s volatility, or how easily it evaporates, and is typically expressed in pounds per square inch (psi) or kilopascals (kPa). The vapor pressure affects the fuel’s ability to vaporize and mix with air in the engine. Proper volatility ensures smooth engine starting, especially in cold weather. In general, lighter hydrocarbons (e.g., butane) increase vapor pressure, while heavier hydrocarbons reduce vapor pressure. Blending components, such as the ethanol, tend to increase vapor pressure, particularly at low blending levels (e.g., 10% ethanol or E10). Hence, the low vapor pressure of the e-naphtha fraction can be improved with additives. Another specification that was not met was the octane number, which was really low for the e-naphtha fraction (40) compared to the gasoline specifications (>95). The octane number could potentially be increased by the use of additives (octane boosters), such as methyl tert-butyl ether (MTBE), and fuel cracking and reforming processes could also optimize the hydrocarbon structure to produce higher-octane fuels. Another possible way to increase the octane number is to blend it with bioethanol, which is characterized by a high octane number (~113) and can be used in internal combustion engines without modification [90].

Regarding the distillation curve of the examined fuels, the accuracy between the cut points of the batch distillation unit that was employed is not considered sufficiently high, so it is likely that some jet boiling range hydrocarbons remained in the naphtha fraction. However, if the fractions were separated with higher accuracy, it is expected that the distillation curve would follow the specifications.

The blend of e-naphtha with the fossil gasoline (90/10 v/v % commercial gasoline/e-naphtha) met almost all gasoline specifications, with the exception of the octane number, which was slighter lower (92.2) compared to the specifications (>95). As discussed previously, the octane number can be improved with the addition of an octane booster, like MTBE or bioethanol. As a consequence, the produced e-naphtha could presumably be utilized as a drop-in fuel in blends with fossil gasoline up to 10 v/v %.

Table 9. Properties of the e-naphtha fraction, the commercial gasoline and their blend (gasoline specifications are also presented).

Properties	Units	Method	EN 228 Gasoline Specifications	Commercial Gasoline	e-Naphtha	Blend (90/10 v/v % Commercial Gasoline/e-Naphtha Fraction)
Density at 288 K	g/ml	EN ISO 12185 [56]	0.7200–0.7750	0.7469	0.7277	0.7456
Vapor pressure at 310.9 K-Mini	kPa	EN-13016-1 [91]	45-60	57.6	11.9	52.9
Total sulfur	ppm-w	EN ISO 20846 [58]	<10	0.5	0.7	0.5
Research octane number, RON		EN ISO 5164 [92]	>95	96.8	<40.0	92.2
Motor octane number, MON		EN ISO 5163 [93]	>85	86.5	<40.0	84.4
Copper strip corros.,3 h-323 K		EN ISO 2160 [63]	Class 1	Class 1a	Class 1a	Class 1a
Existent gum (solvent washed)	mg/100 mL	EN ISO 6246 [94]	<5	3	0	1
Benzene	v/v %	EN ISO 22854 [95]	<1	0.90	0.00	0.82
Aromatics	v/v %	EN ISO 22854 [95]	<35	33.2	3.1	30.6
Olefins	v/v %	EN ISO 22854 [95]	<18	1.1	0.3	1.0
Bio-methanol	v/v %	EN ISO 22854 [95]	<3	0.0	0.0	0.0
Bio-ethanol	v/v %	EN ISO 22854 [95]	<5	0.3	0.0	0.3
Bio-ether > 5C (MTBE-ETBE-TAME)	v/v %	EN ISO 22854 [95]		13.2	0.0	11.9
Oxygen content	w/w %	EN ISO 22854 [95]	<2.7	2.2	0.0	2.0
Appearance		VISUAL	Clear & Bright	Clear & Bright	Clear & Bright	Clear & Bright
Color		VISUAL	Undyed	Undyed	Undyed	Undyed
I.B.P.	K	EN ISO 3405 [70]		307.6	349.3	307.7
F.B.P.	K	EN ISO 3405 [70]	<483	452.6	500.4	464.5
Evaporated at 343.1 K	v/v %	EN ISO 3405 [70]	20–48	33.2	0.0	28.3
Evaporated at 373.1 K	v/v %	EN ISO 3405 [70]	46–71	61.4	4.4	56.2
Evaporated at 423.1 K	v/v %	EN ISO 3405 [70]	>75	91.6	56.9	88.7
Residue	v/v %	EN ISO 3405 [70]	<2	1.0	0.8	1.0

Table 9. Properties of the e-naphtha fraction, the commercial gasoline and their blend (gasoline specifications are also presented).

Properties	Units	Method	EN 228 Gasoline Specifications	Commercial Gasoline	e-Naphtha	Blend (90/10 v/v % Commercial Gasoline/e-Naphtha Fraction)
Density at 288 K	g/ml	EN ISO 12185 [56]	0.7200–0.7750	0.7469	0.7277	0.7456
Vapor pressure at 310.9 K-Mini	kPa	EN-13016-1 [91]	45-60	57.6	11.9	52.9
Total sulfur	ppm-w	EN ISO 20846 [58]	<10	0.5	0.7	0.5
Research octane number, RON		EN ISO 5164 [92]	>95	96.8	<40.0	92.2
Motor octane number, MON		EN ISO 5163 [93]	>85	86.5	<40.0	84.4
Copper strip corros.,3 h-323 K		EN ISO 2160 [63]	Class 1	Class 1a	Class 1a	Class 1a
Existent gum (solvent washed)	mg/100 mL	EN ISO 6246 [94]	<5	3	0	1
Benzene	v/v %	EN ISO 22854 [95]	<1	0.90	0.00	0.82
Aromatics	v/v %	EN ISO 22854 [95]	<35	33.2	3.1	30.6
Olefins	v/v %	EN ISO 22854 [95]	<18	1.1	0.3	1.0
Bio-methanol	v/v %	EN ISO 22854 [95]	<3	0.0	0.0	0.0
Bio-ethanol	v/v %	EN ISO 22854 [95]	<5	0.3	0.0	0.3
Bio-ether > 5C (MTBE-ETBE-TAME)	v/v %	EN ISO 22854 [95]		13.2	0.0	11.9
Oxygen content	w/w %	EN ISO 22854 [95]	<2.7	2.2	0.0	2.0
Appearance		VISUAL	Clear & Bright	Clear & Bright	Clear & Bright	Clear & Bright
Color		VISUAL	Undyed	Undyed	Undyed	Undyed
I.B.P.	K	EN ISO 3405 [70]		307.6	349.3	307.7
F.B.P.	K	EN ISO 3405 [70]	<483	452.6	500.4	464.5
Evaporated at 343.1 K	v/v %	EN ISO 3405 [70]	20–48	33.2	0.0	28.3
Evaporated at 373.1 K	v/v %	EN ISO 3405 [70]	46–71	61.4	4.4	56.2
Evaporated at 423.1 K	v/v %	EN ISO 3405 [70]	>75	91.6	56.9	88.7
Residue	v/v %	EN ISO 3405 [70]	<2	1.0	0.8	1.0

4. Discussion

Hydroprocessing of lipid-based feedstocks for paraffinic diesel and SAF fuels has been demonstrated at the commercial scale; however, the production of naphtha boiling range hydrocarbons has not been commercialized yet. Naphtha boiling range hydrocarbons (340–493 K) are lighter compared to those of jet (493–573 K) and diesel (573–633 K). Therefore, hydroprocessing of lipid-based feedstocks targeting the production of naphtha boiling range hydrocarbons requires more severe operating conditions combined with isomerization, dewaxing and hydrocracking catalysts. Today, studies on the hydroprocessing of lipid-based feedstocks for naphtha boiling range hydrocarbons are limited. As a result, there is a gap in the literature in the current field that this study attempts to fill. Particularly, this study aimed to optimize the hydroprocessing of pretreated WCOs for the production of naphtha boiling range hydrocarbons that could be utilized as a drop-in fuel in commercial gasoline.

The WCOs used in the current research originated from restaurants and households and mainly consisted of triglycerides and fatty acids. A pretreatment process for WCOs was developed, leading to their qualitative improvement prior to the valorization step. A three-stage experimental process was developed to refine WCOs by removing the solids removal, free fatty acids and water content. The quality of the WCOs was notably enhanced, as indicated by the decreases in acidity and water content. In the next step, the refined WCOs were catalytically hydrotreated at a small-scale (TRL-3) continuous flow hydrotreating plant by testing the effect of various operating conditions in order to maximize the naphtha yields, utilizing four different catalysts. A heteroatom removal catalyst followed by a saturation catalyst was combined with an isomerization–dewaxing and a hydrocracking catalyst. It was found that the fuel yields were strongly influenced by the severity of the process, and more specifically, by the temperature, pressure and LHSV. The mild operating conditions promoted diesel production, while stronger conditions led to the production of lighter hydrocarbons. For instance, an increase in temperature from 603 K to 633 K increased the jet selectivity from 11.5% to 22.4%. The influence of LHSV or retention time was even stronger, where a decrease in the LHSV from 1 h⁻¹ to 0.33 h⁻¹ increased naphtha selectivity from 3.6% to 34.3%. In general, the highest naphtha selectivity was achieved at LHSV 0.33 h⁻¹, reaching 34.3%; however, jet and diesel selectivity remained at important levels, at 22.2% and 43.4%, respectively. These findings illustrate that the operating conditions are crucial for such studies and should be chosen based on the targeted products. Under all tested conditions, the conversion rates were high (<96%), but the highest was observed in condition 5 (99%).

Subsequently, the technology was validated and demonstrated at a larger-scale (TRL-5) hydrotreatment plant, targeting the production of higher quantities of organic liquid product for further separation into the different fuel fractions (naphtha, jet and diesel) in order to perform a quality evaluation of the produced e-fuels. The larger-scale unit (TRL-5) consists of two reactors that operate in series. The loading plan from the smaller unit (TRL-3) was adapted to match the setup of the larger-scale unit (TRL-5). The results show that when the HDO and saturation catalysts were loaded in the first reactor and the isomerization with the hydrocracking catalyst in the second reactor, the operation of the smaller (TRL-3) unit was not replicable. The main reason is that during hydrotreatment of refined WCOs in the first reactor, carbon monoxide, carbon dioxide and aqueous phase were formed and poisoned the active sites of the isomerization–dewaxing and hydrocracking catalysts of the second reactor, which led to very low naphtha yields. Therefore, it was decided to perform the hydrotreatment of the refined WCO in a two-step process. In the first step, both reactors were filled only with HDO, HDS, HDN and saturation catalysts in order to remove the heteroatoms from the feed leading in a two-phase liquid product consisting of the organic and the aqueous phase. After the first step, the organic phase was separated from the aqueous phase via gravity, and the reactors were loaded with dewaxing and hydrocracking catalysts to perform the second step. Therefore, neither the aqueous phase nor the gas products from the first step passed through the dewaxing and hydrocracking catalysts. In the end, the resulting product consisted of naphtha, jet and diesel boiling range hydrocarbons. The total organic liquid product from all DOS was collected and fractionated in a batch fractionation plant to separate the three different fractions (naphtha, jet and diesel) for quality evaluation. The findings highlight that the addition of 10 v/v % e-diesel on fossil diesel improved some fuel characteristics of the fossil diesel, such as the cetane index, the flash point, etc., while meeting all fuel specifications for commercial diesel. Regarding the e-jet fraction, the addition of 10 v/v % of e-jet on fossil jet led to a blend that met almost all Jet A1 specifications, with the exception of the freezing point, which was slightly higher (234 K) compared to the specifications (<226 K). However, this can be overcome by adding a lower percentage of e-jet to the blend or with additives that can improve the freezing point. Regarding the e-naphtha fraction, when it was used at 10 v/v % with 90 v/v % fossil gasoline, the blend met almost all gasoline specifications, with the exception of the octane number which was slighter lower (92.2) compared to the gasoline specifications (>95). However, the octane number can be improved through the use of an octane booster like MTBE or bioethanol. As a result, the produced e-naphtha can be easily used as a drop-in fuel in blends with fossil gasoline up to 10 v/v %.

In conclusion, the current manuscript provides novel and important data for industrially relevant applications in the production of advanced e-naphtha, e-jet and e-diesel via hydroprocessing of refined WCOs and lays the foundation for future research aimed at further optimization of the process for high naphtha yields.

5. Conclusions

The current study investigated the technology for electrified fuel production, focusing on e-naphtha fuel via hydrotreatment of pretreated WCOs at an industrially relevant scale. The main findings are summarized below:

• A pretreatment process for WCOs was developed, improving their quality prior to catalytic hydrotreatment in terms of acidity and water content reduction.
• The severity of the hydrotreating process significantly affected the yields of the expected fuels.
• Higher naphtha yields were observed at an operating temperature of 663 K, pressure of 13.78 MPa and liquid hourly space velocity of 0.33 h^−1, leading to 34 wt% naphtha, 23 wt% jet and 42 wt% diesel boiling range hydrocarbons.
• The process was successfully scaled up to the large-scale (TRL-5) hydrotreatment plant.
• The produced fractions were characterized and compared with the fossil fuel standards for diesel, gasoline and Jet A1.
• The blend of e-diesel with fossil diesel (at 10/90 v/v %) meets all specifications for commercial diesel and could be characterized as a high-quality alternative advanced e-diesel fuel.
• The e-jet fraction met the specifications for Jet A1, with the exception of the freezing point, which can be improved with the addition of an extra cyclization step in the process or with the use of freezing point improvers (additives).
• The blend of e-naphtha with fossil gasoline (at 10 v/v %) met almost all specifications, with the exclusion of the octane number; however, this drawback can be overcome with the use of some octane boosters or octane additives.